Combination dehydrocyclodimerization and dehydrogenation process for producing aromatic and olefin products

ABSTRACT

This invention deals with a process for converting aliphatic C 2  -C 6  hydrocarbons into C 6   +  aromatics and C 3   =  /C 4   =  olefins. The process involves combining dehydrocyclodimerization (DHCD) with dehydrogenation. Thus, the feedstream is first sent to a DHCD zone which produces an effluent stream which contains C 6   +  aromatics along with C 1  -C 5  hydrocarbons. This effluent stream is separated into a stream containing C 1  -C 4  hydrocarbons and one containing C 6   +  aromatics. The C 1  -C 4  containing stream is flowed to a dehydrogenation zone to produce C 3   =  /C 4   =  olefins.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a Continuation-in-Part of prior application U.S.Ser. No. 08/167,299 filed on Dec. 16, 1993, U.S. Pat. No. 5,401,893.

FIELD OF THE INVENTION

This invention relates to a hydrocarbon conversion process.Specifically, a light aliphatic hydrocarbon is first converted toaromatics plus hydrogen. The hydrogen and unconverted hydrocarbons arenext flowed to a dehydrogenation zone where the hydrocarbons areconverted to olefins.

BACKGROUND OF THE INVENTION

Dehydrocyclodimerization (DHCD) is a process in which aliphatichydrocarbons containing from 2 to 6 carbon atoms per molecule arereacted over a catalyst to produce a high yield of aromatics andhydrogen. This process is well known and is described in detail in U.S.Pat. Nos. 4,654,455 and 4,746,763 which are incorporated by reference.Typically, the dehydrocyclodimerization reaction is carried out attemperatures in excess of 500° C., using dual functional catalystscontaining acidic and dehydrogenation components. The acidic function isusually provided by a zeolite which promotes the oligomerization andaromatization reactions, while a non-noble metal component promotes thedehydrogenation function.

Since the product stream from the dehydrocyclodimerization processcontains a mixture of compounds, it must undergo several separationsteps in order to obtain usable products. An initial fractionation willseparate the C₆ ⁺ products from uncondensed materials which include fuelgas, hydrogen and unreacted hydrocarbons. The uncondensed material iscompressed and sent to a gas recovery section where hydrogen and fuelgas are separated from the unconverted hydrocarbons which are recycledto the dehydrocyclodimerization zone.

Significant fixed and operating costs are associated with the productcompressor and gas recovery equipment. The high compression ratiorequired to condense the materials accounts for the substantial capitaland operating costs. Therefore, it would be desirable to eliminate thehigh costs associated with the compressor and gas recovery equipment.

Applicants have found a solution to this problem which involvesdehydrogenation. The uncondensed materials from the product fractionatorare flowed to a dehydrogenation zone where the unreacted hydrocarbons(C₃ /C₄) are converted to olefins. Since the operating pressure of thedehydrogenation zone is lower than that of the DHCD zone, there is noneed for a compressor. The effluent stream from the dehydrogenation zoneis separated into a hydrogen stream, a C₁ /C₂ fuel gas stream, a C₃ ⁼/C₄ ⁼ olefin stream and an unreacted hydrocarbon stream (compression isrequired for this separation). The unreacted stream is recycled to theDHCD zone, while the olefin stream is collected and the other streamsare vented.

SUMMARY OF THE INVENTION

As stated, the present invention relates to a process for converting C₂-C₆ aliphatic hydrocarbons to aromatics and C₃ ⁼ /C₄ ⁼ olefins.Accordingly, one embodiment of the invention is a process for producingaromatic and olefinic hydrocarbons comprising:

a) flowing a hydrocarbon feedstream containing C₂ -C₆ aliphatichydrocarbons into a dehydrocyclodimerization zone where said feedstreamis contacted with a bed of a solid dehydrocyclodimerization catalyst atdehydrocyclodimerization conditions, thereby producing an effluentstream containing C₆ ⁺ aromatics, hydrogen and C₁ -C₄ aliphatichydrocarbons;

b) flowing the effluent stream to a first separation zone and separatingthe stream to obtain a first bottoms stream containing C₆ ⁺ aromaticsand a first overhead stream containing hydrogen and C₁ -C₄ aliphatichydrocarbons;

c) flowing the first overhead stream to a dehydrogenation zone andcontacting said stream with a dehydrogenation catalyst in the presenceof hydrogen at dehydrogenation conditions, thereby producing a secondeffluent stream containing C₃ ⁼ /C₄ ⁼ olefins, hydrogen and unreacted C₁-C₄ aliphatic hydrocarbons;

d) flowing the second effluent stream to a gas recovery zone andseparating the second effluent stream into a hydrogen stream, a fuel gas(C₁ /C₂) stream, a C₃ ⁼ /C₄ ⁼ and a C₃ -C₄ recycle aliphatic stream;

e) recycling at least a portion of the C₃ -C₄ aliphatic stream to thedehydrogenation zone of step (c).

This and other objects and embodiments of this invention will becomemore apparent after a detailed description of the invention.

BRIEF DESCRIPTION OF THE DRAWING

The FIGURE is a simplified process flow diagram of one embodiment of theinvention showing the formation of aromatic and olefinic (C₃ ⁼ /C₄ ⁼)hydrocarbons from a C₂ -C₆ aliphatic hydrocarbon stream.

DETAILED DESCRIPTION OF THE INVENTION

As stated, this invention relates to a process for preparing botharomatic and light olefinic (C₃ ⁼ /C₄ ⁼) hydrocarbons from a lightaliphatic hydrocarbon stream. The process uses adehydrocyclodimerization zone and a dehydrogenation zone. The feedstreamto the present process contains light aliphatic hydrocarbons having from2 to 6 carbon atoms per molecule. The feedstream may contain a singlecompound or a mixture of two or more of these compounds. Preferredcompounds are propane and butanes. It is also preferred that theconcentration of C₅ and C₆ hydrocarbons in the feedstream be held to apractical level, preferably below 20 mole percent.

Regardless of the composition of the feedstream, the feedstream isflowed into a dehydrocyclodimerization (DHCD) zone which converts asignificant portion of the aliphatic feedstream into aromatichydrocarbons, i.e., C₆ ⁺ aromatics. The majority of the C₆ ⁺ producthydrocarbons are benzene, toluene and the various xylene isomers with asmall amount (about 8%) of C₉ ⁺ aromatics. Hydrogen is also a product ofthe process. Since conversion of the aliphatic feedstream to aromaticproducts is not one hundred percent, the product effluent stream willalso contain unreacted feedstock along with C₁ /C₂ hydrocarbons. Theconfiguration of the reaction zone and the composition of the catalystemployed in the reaction zone are well known in the art and aredescribed here for completeness.

Usually the reaction zone consists of a moving bed radial flowmultistage reactor as described, for example, in U.S. Pat. Nos.4,110,081 and 4,403,909 which are incorporated by reference. Thesepatents also describe catalyst regeneration systems and various aspectsof moving catalyst bed operations and equipment. A preferred moving bedreactor system employs a spherical catalyst having a diameter betweenabout 0.4 millimeters and 3.2 millimeters. The catalyst preferablycomprises a zeolitic material, a metallic component and a binder. U.S.Pat. Nos. 4,654,455, 4,746,763 and 5,169,812, which are incorporated byreference, describe DHCD catalysts and methods of preparing them. Abrief description of these catalysts will be presented.

The zeolites which may be used are any of those which have a Si:Al ratiogreater than about 10 and preferably greater than 20 and a pore diameterof about 5 to 6 Angstroms. Specific examples of zeolites which can beused are the ZSM family of zeolites. Included among this ZSM family areZSM-5, ZSM-8, ZSM-11, ZSM-12 and ZSM-35. The preparation of theseZSM-type zeolites is well known in the art and generally are prepared bycrystallizing a mixture containing an alumina source, a silica source,an alkali metal source, water and a tetraalkyl ammonium compound or itsprecursor. The amount of zeolite present in the catalyst can varyconsiderably but usually is present in an amount from about 30 to about90 weight percent and preferably from about 50 to about 70 weightpercent of the catalyst.

A second component of these catalysts is a phosphorus containing alumina(hereinafter referred to as aluminum phosphate) component. Thephosphorus may be incorporated with the alumina in any acceptable mannerknown in the art. One preferred method of preparing this aluminumphosphate is that described in U.S. Pat. No. 4,629,717 which isincorporated by reference. The technique described in the '717 patentinvolves the gellation of a hydrosol of alumina which contains aphosphorus compound using the well-known oil drop method. Generally thistechnique involves preparing a hydrosol by digesting aluminum in aqueoushydrochloric acid at reflux temperatures of about 80° to 105° C. Theratio of aluminum to chloride in the sol ranges from about 0.7:1 toabout 1.5:1 weight ratio. A phosphorus compound is now added to the sol.Preferred phosphorus compounds are phosphoric acid, phosphorous acid andammonium phosphate. The relative amount of phosphorus and aluminumexpressed in molar ratios ranges from about 1:1 to 1:100 on an elementalbasis.

The resulting aluminum phosphate hydrosol mixture is now gelled. Onemethod of gelling this mixture involves combining a gelling agent withthe mixture and then dispersing the resultant combined mixture into anoil bath or tower which has been heated to elevated temperatures suchthat gellation occurs with the formation of spheroidal particles. Thegelling agents which may be used in this process are hexamethylenetetraamine, urea or mixtures thereof. The gelling agents release ammoniaat the elevated temperatures which sets or converts the hydrosol spheresinto hydrogel spheres. The spheres are then continuously withdrawn fromthe oil bath and typically subjected to specific aging and dryingtreatments in oil and in ammoniacal solution to further improve theirphysical characteristics. The resulting aged and gelled particles arethen washed and dried at a relatively low temperature of about 93° C. toabout 149° C. (200°-300° F.) and subjected to a calcination procedure ata temperature of about 450° C. to about 703° C. (850°-1300° F.) for aperiod of about 1 to about 20 hours. The amount of phosphorus containingalumina component present (as the oxide) in the catalyst can range fromabout 10 to about 70 weight percent and preferably from about 30 toabout 50 weight percent.

The zeolite and aluminum phosphate binder are mixed and formed intoparticles by means well known in the art such as gellation, pilling,nodulizing, marumerizing, spray drying, extrusion or any combination ofthese techniques. A preferred method of preparing the zeolite/aluminumphosphate support involves adding the zeolite either to an alumina solor a phosphorus compound, forming a mixture of the aluminasol/zeolite/phosphorus compound which is now formed into particles byemploying the oil drop method described above. The particles arecalcined as described above to give a support.

Another component of these catalysts is a gallium component. The galliumcomponent may be deposited onto the support in any suitable manner knownto the art which results in a uniform dispersion of the gallium. Usuallythe gallium is deposited onto the support by impregnating the supportwith a salt of the gallium metal. The particles are impregnated with agallium salt selected from the group consisting of gallium nitrate,gallium chloride, gallium bromide, gallium hydroxide, gallium acetate,etc. The amount of gallium which is deposited onto the support variesfrom about 0.1 to about 5 weight percent of the finished catalystexpressed as the metal.

The gallium compound may be impregnated onto the support particles byany technique well known in the art such as dipping the catalyst into asolution of the metal compounds or spraying the solution onto thesupport. One preferred method of preparation involves the use of a steamjacketed rotary dryer. The support particles are immersed in theimpregnating solution contained in the dryer and the support particlesare tumbled therein by the rotating motion of the dryer. Evaporation ofthe solution in contact with the tumbling support is expedited byapplying steam to the dryer jacket. After the particles are completelydry, they are heated under a hydrogen atmosphere at a temperature ofabout 500° to about 700° C. for a time of about 1 to about 15 hours.Although a pure hydrogen atmosphere is preferred to reduce and dispersethe gallium, the hydrogen may be diluted with nitrogen. Alternatively,it is envisioned that the reduction and dispersion can be done in situin the actual reactor vessel used for dehydrocyclodimerization by usingeither pure hydrogen or a mixture of hydrogen and hydrocarbons. Next thehydrogen treated particles are heated in air and steam at a temperatureof about 400° to about 700° C. for a time of about 1 to about 10 hours.The amount of steam present in the air varies from about 1 to about 40percent.

A particularly preferred catalyst is one described in U.S. Pat. No.5,169,812. This reference describes a gallium/aluminum phosphate/zeolitecatalyst which has been treated with an aqueous solution of a weaklyacidic ammonium salt or a dilute acid solution, e.g., ammonium chlorideor hydrochloric acid.

The dehydrocyclodimerization conditions which are employed in thereaction zone will vary depending on such factors as feedstockcomposition and desired conversion. A desired range of conditions forthe dehydrocyclodimerization of C₂ -C₆ aliphatic hydrocarbons toaromatics include a temperature from about 350° C. to about 650° C., apressure from about 100 kPa to about 2,020 kPa and a liquid hourly spacevelocity from about 0.2 to about 5 hrs⁻¹. The preferred processconditions are a temperature in the range from about 400° to about 550°C., a pressure in the range from about 200 to about 1,015 kPa and aliquid hourly space velocity of between 1.0 to 4.0 hrs⁻¹. It isunderstood that as the average carbon number of the feedstreamincreases, a temperature in the lower end of the temperature range isrequired for optimum performance and conversely as the average carbonnumber of the feed decreases, the higher the required temperature.

The effluent stream from the DHCD zone which contains C₆ ⁺ aromatics,hydrogen, unreacted feed and C₁ /C₂ is separated into a streamcontaining C₆ ⁺ aromatics and one containing the other components. Thestream containing hydrogen, unreacted feedstock and C₁ /C₂ hydrocarbonsis now fed to a dehydrogenation zone. Dehydrogenation reactions zonesare well known in the art as exemplified by U.S. Pat. Nos. 4,447,653 and4,868,342 which are incorporated by reference. The dehydrogenation zonelike the DHCD zone comprises at least one radial flow reactor. Thedehydrogenation reaction is a highly endothermic reaction which istypically effected at low (near atmospheric) pressure conditions. Theprecise dehydrogenation temperature and pressure employed in thedehydrogenation reaction zone will depend on a variety of factors suchas the composition of the paraffinic hydrocarbon feedstock, the activityof the selected catalyst, and the hydrocarbon conversion rate. Ingeneral, dehydrogenation conditions include a pressure of from about 0to about 3,500 kPa and a temperature of from about 480° C. (900° F.) toabout 760° C. (1400° F.). The hydrocarbons are charged to the reactionzone and contacted with the catalyst contained therein at a liquidhourly space velocity of from about 1 to about 10. Hydrogen is suitablyadmixed with the hydrocarbon feedstock in a mole ratio of from about 0.1to about 10. Preferred dehydrogenation conditions, particularly withrespect to C₃ -C₄ paraffinic hydrocarbon feedstocks, include a pressureof from about 0 to about 2,000 kPa and a temperature of from about 540°C. (1000° F.) to about 705° C. (1300° F.), a liquid hourly spacevelocity of from about 1 to about 5 hr⁻¹, and a hydrogen/hydrocarbonmole ratio of about 0.5 to about 2.

The dehydrogenation zone of this invention may use any suitabledehydrogenation catalyst. Generally, the preferred catalyst comprises aplatinum group metal, an alkali metal component, and a porous inorganiccarrier material. The catalyst may also contain one or more modifiermetal which advantageously improve the performance of the catalyst. Itis preferable that the porous carrier material of the dehydrogenationcatalyst be an absorptive high surface area support having a surfacearea of about 25 to about 500 m² /g. The porous carrier material shouldbe relatively refractory to the conditions utilized in the reaction zoneand may be chosen from those carrier materials which have traditionallybeen utilized in dual function hydrocarbon conversion catalysts. Aporous carrier material may, therefore, be chosen from an activatedcarbon, coke or charcoal, silica or silica gel, clays and silicatesincluding those synthetically prepared and naturally occurring, whichmay or may not be acid-treated as, for example, attapulgus clay,diatomaceous earth, kieselguhr, bauxite; refractory inorganic oxidessuch as alumina, titanium dioxide, zirconium dioxides, magnesia, silicaalumina, alumina boria, etc. or a combination of one or more of thesematerials. The preferred porous carrier material is a refractoryinorganic oxide, with the best results being obtained with an aluminacarrier material. The aluminas, such as gamma alumina, give the bestresults in general. The preferred catalyst will have a gamma aluminacarrier which is in the form of spherical particles having relativelysmall diameters on the order of about 1/16 inch.

The preferred dehydrogenation catalyst also contains a platinum groupmetal. Of the platinum group metals, which include palladium, rhodium,ruthenium, osmium and iridium, the use of platinum is preferred. Theplatinum group metal may exist within the final catalyst composite as acompound such as an oxide, sulfide, halide, oxysulfide, etc., or anelemental metal or in combination with one or more other ingredients ofthe catalyst. It is believed that the best results are obtained whensubstantially all the platinum group metal exists in the elementalstate. The platinum group metal generally comprises from about 0.01 toabout 2 wt. % of the final catalytic composite, calculated on anelemental basis. It is preferred that the platinum content of thecatalyst be between about 0.1 and 1 wt. %. The preferred platinum groupmetal is platinum, with palladium being the next preferred metal. Theplatinum group metal may be incorporated into the catalyst composite inany suitable manner such as by coprecipitation or coagulation with thepreferred carrier material, or by ion exchange or impregnation of thecarrier material. The preferred method of preparing the catalystnormally involves the utilization of a water-soluble, decomposablecompound of a platinum group metal to impregnate the calcined carriermaterial. For example, the platinum group component may be added to thesupport by commingling the support with an aqueous solution ofchloroplatinic or chloropalladic acid. An acid such as hydrogen chlorideis generally added to the impregnation solution to aid in thedistribution of the platinum group component throughout the carriermaterial.

Additionally, the preferred catalyst contains an alkali metal componentchosen from cesium, rubidium, potassium, sodium and lithium. Thepreferred alkali metal is normally either potassium or lithium,depending on the feed hydrocarbon. The concentration of the alkali metalmay range from about 0.1 to 5 wt. %, but is preferably between 1 andabout 4 wt. % calculated on an elemental basis. This component may beadded to the catalyst by the methods described above as a separate stepor simultaneously with the solution of another component. With somealkali metals, it may be necessary to limit the halogen content to lessthan 0.5 wt. % and preferably less than 0.1 wt. %, while others may havehigher halogen content.

As noted previously, the dehydrogenation catalyst may also contain amodifier metal. One such preferred promoter metal is tin. The tincomponent should constitute about 0.01 to about 1 wt. % tin. It ispreferred that the atomic ratio of tin to platinum be between 1:1 andabout 6:1. The tin component may be incorporated into the catalyticcomposite in any suitable manner known to effectively disperse thiscomponent onto the carrier material.

The FIGURE illustrates one embodiment of the invention. Those skilled inthe art will recognize that this process flow diagram has beensimplified by the elimination of many pieces of process equipmentincluding heat exchangers, process control systems, pumps, fractionationcolumn overhead and reboiler systems, etc. which are not necessary to anunderstanding of the process. It may also be readily discerned that theprocess flow presented in the drawings may be modified in many aspectswithout departing from the basis overall concept of the invention.Referring now to the FIGURE, a feedstream of C₂ -C₆ hydrocarbons andpreferably C₃ -C₄ is flowed via line 1 into reaction zone 2. As stated,reaction zone 2 is a dehydrocyclodimerization (DHCD) zone where the C₂-C₆ hydrocarbon feedstream is contacted with a solid catalyst under DHCDconditions. The effluent stream, which contains C₆ ⁺ aromatics,hydrogen, C₁ /C₂ and unreacted feedstock, is removed via line 3 andflowed to separation zone 4.

Separation zone 4 is operated at conditions sufficient to separate theeffluent stream into a first bottoms stream containing C₅ ⁺ hydrocarbonsand a first overhead stream containing hydrogen and C₁ -C₄ aliphatichydrocarbons. The first bottoms stream contains aromatic and a smallamount of non-aromatic hydrocarbons, (C₅ ⁺), with the majority of thestream composed of C₆ ⁺ aromatics, especially benzene, toluene andxylenes. The first bottoms stream is removed via line 5 and is eithercollected as is or is further processed to separate the C₆ ⁺ aromaticsinto benzene, toluene and xylenes products.

The first overhead stream is removed from zone 4 via line 6 and flowedto dehydrogenation zone 7. In zone 7 the stream is contacted with adehydrogenation catalyst which converts the C₃ /C₄ saturatedhydrocarbons to olefins (C₃ ⁼ /C₄ ⁼). This second effluent stream isremoved via line 8 and flowed to gas recovery zone 9. This zone isoperated at conditions to separate the components of the second effluentstream into a hydrogen stream which is removed via line 10, a fuel gas(C₁ /C₂) stream which is removed via line 11, a C₃ ⁼ /C₄ ⁼ productstream which is removed via line 12 and a C₃ -C₄ recycle aliphaticstream which is removed via line 13. The C₃ -C₄ aliphatic stream is nowrecycled in one of several ways. A portion of the C₃ -C₄ aliphaticstream is flowed via line 13 to line 14 into dehydrogenation zone 7,while another portion is flowed via line 13 into line 1 and then intoDHCD zone 2. Alternatively, substantially all of the C₃ -C₄ aliphaticstream is flowed via line 13 into line 14 and into dehydrogenation zone7. Finally, substantially all of the C₃ -C₄ aliphatic stream is flowedvia line 13 into line 1 and recycled into DHCD zone 2. A compressor isrequired as part of zone 9 in order to effectively separate the variouscomponents.

As stated, the instant process presents several advantages over the art.First, capital costs are significantly reduced owing to the fact thatcompression of the product stream from the dehydrocyclodimerization zoneis not necessary. Second, the instant process can provide the totalfeedstock requirements for a downstream aromatics derivate application.Finally, optimal feedstock utilization is obtained by elimination ofseparation inefficiencies.

We claim as our invention:
 1. A process for producing aromatic andolefinic hydrocarbons comprising:a) flowing a hydrocarbon feedstreamcontaining C₂ -C₆ aliphatic hydrocarbons into a dehydrocyclodimerizationzone where said feedstream is contacted with a bed of a soliddehydrocyclodimerization catalyst at dehydrocyclodimerizationconditions, thereby producing an effluent stream containing C₆ ⁺aromatics, hydrogen and C₁ -C₄ aliphatic hydrocarbons; b) flowing theeffluent stream to a first separation zone and separating the stream toobtain a first bottoms stream containing C₆ ⁺ aromatics and a firstoverhead stream containing hydrogen and C₁ -C₄ aliphatic hydrocarbons;c) flowing the first overhead stream to a dehydrogenation zone andcontacting said stream with a dehydrogenation catalyst in the presenceof hydrogen at dehydrogenation conditions, thereby producing a secondeffluent stream containing C₃ ⁼ /C₄ ⁼ olefins, hydrogen and unreacted C₁-C₄ aliphatic hydrocarbons; d) flowing the second effluent stream to agas recovery zone and separating the second effluent stream into ahydrogen stream, a fuel gas (C₁ /C₂) stream, a C₃ ⁼ /C₄ ⁼ stream and aC₃ -C₄ recycle aliphatic stream; e) recycling at least a portion of theC₃ -C₄ aliphatic stream to the dehydrogenation zone of step (c).
 2. Theprocess of claim 1 characterized in that at least a portion of the C₃-C₄ aliphatic stream is recycled to the dehydrocyclodimerization zone ofstep (a).
 3. The process of claim 1 characterized in that substantiallyall of the C₃ -C₄ aliphatic stream is recycled to the dehydrogenationzone of step (c).
 4. The process of claim 1 where thedehydrocyclodimerization conditions include a temperature of about 350°C. to about 650° C., a pressure of about 100 kPa to about 2,020 kPa anda liquid hourly space velocity of about 0.2 to about 5 hr⁻¹.
 5. Theprocess of claim 1 where the dehydrocyclodimerization catalyst comprisesa zeolite component, a gallium component and an aluminum phosphatebinder.
 6. The process of claim 1 where the dehydrogenation conditionsinclude a temperature of about 480° C. to about 760° C., a pressure ofabout 0 to about 3,500 kPa, a liquid hourly space velocity of about 1 toabout 5 hr⁻¹ and a hydrogen to hydrocarbon ratio of about 0.1 to about10.
 7. The process of claim 1 where the dehydrogenation catalystcomprises a platinum group metal dispersed on a porous inorganic carriermaterial.
 8. The process of claim 7 where the platinum group metal isplatinum and is present in an amount from about 0.01 to about 2 weightpercent of the catalyst.
 9. The process of claim 7 further characterizedin that the dehydrogenation catalyst contains an alkali metal component.10. The process of claim 9 where the alkali metal component ispotassium.
 11. The process of claim 7 further characterized in that thedehydrogenation catalyst contains a modifier metal.
 12. The process ofclaim 11 where the modifier metal is tin.